Production of high alcohols by improved oxo process



vMay l2, 1953. w. RUssUM ETAL 2,638,487

PRODUCTION OF HIGH ALCOHOLS BY IMPROVED OXO PROCESS Filed Sept. 13, 1951 2 Sheets-Sheet 1 m E o 'C 03 o o g s; La o a g m m .E E E (D l E N (n 2 O c N 8 i N uoiv/vo/ovad 2 ATTORNEY L. WQ RUSSUM EVAL May 12, 1953 PRODUCTIONVOF HIGHl ALCOHOLS BY IMPROVED OXO PROCESS Filed sept. 1s, 1951 `2 sheets-sheet 2 Patented May 12, 1953 PRODUCTION OF HIGH ALCOHOLS BY IMPROVED OXO PROCESS Leonard W. Bussum, Highland, and Robert J. Hengstebeck, Valparaiso, Ind., assignors to Standard Oil Company, Chicago, Ill., a corporation of Indiana Application September 13, 1951,'Serial'No. 246,368

11y Glaims.

This invention relates to improvement in production of high boiling alcohols and it pertains more particularly to improved methods and means for producing high boiling alcohols which are substantially free from aldehydes and from other impurities which have heretofore contaminated such alcohols when produced by the socalled oxo process. This application is a continuation-in-part of our parent application, Serial Number 20,786, filed April 13, 1948, and now abandoned; certain of the subject matter herein disclosed is being claimed in co-pending companion application, Serial Number 246,433 in the name of William J. Cerveny, filed September 13, 1951, as a continuation-impart of Serial Number 20,753, filed April 13, 1948, noW abandoned.

In the oxo process an olefin, such for example as a heptene or an octene, is reacted with carbon monoxide and hydrogen at high pressure in the presence of a carbonyl-forming metal catalyst such as cobalt to yield an aldehyde with one more carbon atom per molecule. This reaction is referred to as oxolation and it is accompanied by side reactions which include la certain amount of hydrogenation to form alcohols, aldehyde polymerization and formation of relatively high boiling products of indeterminate composition. The second step lof the oxo process is the' hydrogenation of products produced in the iirst step primarily to convert the aldehyde into alcohol' after which the hydrogenated products are fractionated into an alcohol fraction (nonyl alcohol When the charged olefin is an octene), a. lower boiling hydrocarbon fraction and a higher boiling fraction. An object of the invention is to provide improvements in the oxo process whereby product degradation is reduced to a minimum and whereby maximum yieldsof high quality products may be obtained with minimum investment and 0perating costs. Another object is to provide irnproved methods and means of catalyst utilization, recovery and reuse. Still another object is toincreaser the effectiveness of hydrogen utilization.

A further object yof the invention is to provide methods and means for controlling reaction conditions in the oxolation and hydrogenation steps. Another object is to provide improved correlation between gas production and utilization portions of the process. Other objects will be apparent as the detailed description of the invention proceeds.

To accomplish the above objects sources of Hz-CO and H2 are required. Thus a normally gaseous hydrocarbon may be reacted with steam in a multiple reformer system, the products 0f the first reformer being passed through a converter with additional steam to produce a gas consisting essentially of hydrogen and carbon dioxide. This carbon dioxide may be separated and reacted With hydrocarbons and steam 'in a second reformer to produce carbon monoxide and hydrogen in a ratio of about 1:1 together with carbon dioxide which is separated and recycled. The hydrogen produced in the first reformer and converter, after removal of CO2, Will usually contain about .5% to 2% of carbon monoxide; such carbon monoxide may be removed by methanation or other known means to a level rbelow 0.1%.

The 1:1 hydrogen-carbon monoxide mixture (ratios as high as 1.5:1 may be used) is passed through an oxolation reaction Zone together with an aliphatic olefin containing 3 to 15 or more carbon atoms per molecule, for example a mixture of heptenes or octenes, in the presence of an oxolation catalyst such asvcobalt under conditions t-o eiiiect substantial oxo-lation, i. e. conversion of octene to nonyl aldehyde. Where the charge is a mixture of olens obtained by polymerization of a mixture of n-butenes and isobutylene, the

olens which do not react are generally characterized by highest octane numbers and are the mostl valuable components for motor fuel. Oxolation may be effected by operating at Z500-4000, i. e. about 3000 p. s. i. g. at a temperature of about 250 to 400 F., e. g. about 330 F., with a liquid space velocity (volumes of fresh liquid per hour per volume of reactor space) of about 0.15 to 1.5, e. g. about 0.5, employing about 0.01 to 0.2, e. g. about 0.1 weight per cent 4catalyst as cobalt (supplied as an oil-soluble cobalt salt such as a cobalt naphthenate solution) and about 20 to 50 or more, preferably about 25 to 40 cubic feet of hydrogencarbon monoxide gas (in addition to recycle gas) per gallon Eof olefin charged. For oxolaton of butylene the amount of hydrogen-carbon' monoxide gas may be up to about 80, and for oxollation of C15 olens as low as about 15 cubic feet per gallon. The preferred Ioperating conditions above dened result in a total olefin conversion of about 50% (S5-75%)', an aldehyde to alcohol ratio of about 3.1 (1:1 to 5:1), an aldehyde-plusalcohol to bottoms or polymer ratio of about 3:1 (2:1 to 17:1) and olefin saturation of aboutr 7% (0 to 10%). Oxolation temperaturey is stabilized by the large liquid content of the reactor and' may be controlled in any known manner, preferably by cooling reactor eiiluent, separating and recycling product liquid at spaced vertical points in the reactor and recycling gas to the base thereof.

The liquid product of .the oxalation step, after gases are released therefrom and after catalyst has been removed therefrom by acid and water washing, may next be subjected to a first hydrogenation step under conditions for converting most of the nonyl aldehydes to nonyl alcohols with minimum conversion of alcohols .to hydrocarbons. This hydrogenation may be effected by trickling the oXolation product over supported cobalt or copper chromite catalyst at a temperature of about 350 to 550 F., e. g. about 450 F. and a pressure of about 500 .to 3000 p. s. i. g., e. g. about 850 p. s. i. g., in the presence of hydrogen which may be from the final hydrogenation step. To remove the heat of hydrogenation a substantial part of the hydrogenated product is cooled and recycled, the recycle ratio in this case being about 1:1 to 3:1. A large recycle ratio is a, safeguard against undue temperature rise in the event the oleilns undergo hydrogenation.

Under preferred operating conditions the oxolation step may effect sufficient hydrogenation to avoid the necessity of employing a -first hydrogenation step. When the catalyst is removed from the liquid oXolation products by acid wash and the acid is thoroughly removed by water or dilute caustic wash, these products may be steam distilled under reduced pressure at short contact time to remove polymer as bottoms, to remove overhead, if desired, at least a part of the unconverted and saturated olens and to provide as a charge for final hydrogenation a heart cut of C8 aldehyde and alcohol.

When a iirst hydrogenation step is employed, the oXolation product stream, after acid washing to remove cobalt and a rst hydrogenation to increase alcohol content, may be fractionated to remove most of the unreacted hydrocarbons (chiefly codimer and particularly olefins wherein the double bond is between tertiary carbon atoms) and to remove all materials lower boiling than the desired alcohols but it is particularly important to effect removal of materials higher boiling than the desired alcohols. The alcohol product thus obtained may contain about .5 to 3% of nonyl aldehydes. Such impure nonyl alcohol is subjected to a final hydro- .genation step by trickling it over a hydrogenation catalyst such as copper chromite or a supported cobalt catalyst at a temperature of about 300 to 500 F. or more, and a pressure of about 500 to about 3000, in the presence of the hydrogen produced as hereinabove described. A large excess of hydrogen may be introduced into the final hydrogenation reactor and under the described conditions substantially all of the aldehydes are converted into alcohol without appreciably reducing the alcohol to hydrocarbons. The unconsumed hydrogen from the nal hydrogenation step may thereafter be employed in the rst hydrogenation step. The product produced by the final hydrogenation step is an alcohol of high purity which contains substantially less than .5% of aldehyde and which is substantially free from discoloration or color forming materials. At some sacrifice in product yield, the second hydrogenation step may be omitted and the fractionation may be effected at such low pressure as to give a product of good color and low aldehyde content.

When no first hydrogenation step is required (except the hydrogenation which is inherently effected in the oxolation reactor), the liquid products from the oxolation reactor are freed from catalyst by washing with dilute sulfuric acid followed by water washing and the sub-r stantially neutral organic liquid thus obtained is steam distilled under reduced pressure with a short contact time to eliminate polymers, i. e. materials higher boiling than the desired. alcohols, as a bottoms fraction and preferably to eliminate unreacted and saturated olefins as an overhead fraction so that the iinal hydrogenation step, as in the other example, is effected with an aldehyde-alcohol charging stock which is free from materials higher boiling than the alcohol and from which at least a part of the oleiins has been removed. The steam distillation may be effected in the presence or absence of a buffer and should be effected with a preheat temperature and tower bottom temperature sufficiently low and with sufciently short contact time and reduction in pressure to enable the fractionation to be performed without any serious product degradation. An auxiliary heat carrier and/or stripping medium such as recycled olefins may be added tothe stream in the-place of or in addition to steam before the stream is preheated and introduced into the distillation tower.

An important feature of our invention is the recovery and reuse of oxolation catalyst. The oxolation product is washed with sulfuricv acid of about 10% concentration to convert all cobalt metal and cobalt compounds contained therein into cobalt sulfate which is removed as acid extract of about 5% acid concentration. The acid extract is then treated with dilute sodium hydroxide and naphthenic acid in the presence of a liquid hydrocarbon solvent (which may be a portion of the hydrocarbon charged to the oXolation step). Cobalt sulfate is thus converted into cobalt naphthenate and this reaction is substantially quantitative because the precipitated cobalt naphthenate is constantly removed from the aqueous phase by solutionin the hydrocarbon. Aqueous sodium sulfate .and any unreacted caustic is then settled from the resulting cobalt naphthenate solution and the solution may be washed with water and then recycled to the oxolation step.

Temperature control features are important in the` oxolation and hydrogenation steps. Oxolation temperature may be controlled by cooling oxolation effluent at substantially oxolation pressure, separating gases released from the cooled products and recycling said cooled gases to the base of the reactor and liquid products at higher spaced vertical elevations therein. At the high pressures employed this method of temperature control is remarkably effective. Temperature control in the first hydrogenation step is effected by recycling hydrogenated liquid after removing sulcient gas therefrom so that the pump returning the liquid will not become gas bound. The recycle of the fraction of the hydrogenated product which is higher boiling than oleins but lower boiling than alcohol is particularly advantagous as hereinafter described. Temperature control in the final hydrogenation step is easily attained through the temperature of the feed when the heat release is relatively small. Large volumes of hydrogen may be used in this step of the process. When the acid-wash and waterwashed oxolation products are steam distilled and subsequent hydrogenation is limited to the fraction containing the desired number of carbon atoms per molecule, temperature control may be effected by recycling hydrogenated product to the reactor and also by recycling cooled .hydrogen at spaced `points therein.

By recycling hydrogen-carbon monoxide gas released from oxolation products to the absorber from which the hydrogen-carbon monoxide streamis originally obtained, carbon dioxide may be purged landv maximum utilization of the hydrogen-carbon monoxide gas can bel effected. Other features of th'e invention will be apparent :from the following ydetailed descriptionv of a specific example thereof.

The invention will be more clearly understood from the following detailed description vread in conjunction with the accompanying drawings in which:

Fig. 1 isa schematic now diagram of a commercial plant for producing nonyl alcohol, and' Fig. 2 is a vschematic flow diagram of a ccmvmercial plant for producing octyl alcohol.

In order toproduce the hydrogen-carbon monoxide 1:1 gas mixture and the hydrogen for the hydrogenation steps-a multiple gas reformer System is employed. A mixture of steam from line I and hydrocarbon gas (which has been previously scrubbed withA caustic solution and washed with filtered water) from line II- is passed through rst lreformer I2 wherein it is heated to a temperature of about 1500 to 1600" F. and contacted with a known reformer catalyst such as. nickel promotedwith ceria or magnesia (see U. S. 1,904,592) to obtain a product consisting chiefly of H2, C`O= and CO2. rihese products are `cooled to about 80G-900i F. and introduced by line I3 with additional steam from line I4 into converter I5 wherein the mixture iscontacted with a conversion catalyst (such as iron oxide, which may be promoted with CaO', MgO,

.CrzOs and/or other known promoters) for converting. the COl and steam to CO2 and H2. The

Vproducts from the converter, after condensing and separating out water, are introduced by line IE andcompressor I'I into CO2 absorber IBwherein the CO2 is scrubbed out with a lean MEA (monoethanolamine) solution at about 100 F.

and 235 p. s. i. g. so that the gas discharged through line I9 consists essentially of hydrogen although it usually contains about 1 to 2% of carbon monoxide. Since it is desirable to keep the C'O content lower than 0.1%, the CO is preferably removed by any known means (not shown) such as byy catalytic conversion to methane. The rich MEA solution is introduced by line to the upper part of CO2 activator tower ZI wherein the CO2', is stripped out` of the solution, the lean solution being returned by line 22, pump 23, coolers (not shown) and-line 2t to absorber tcwer I8.

The CO2 from the top -of tower 2'I together With hydrocarbon gas from lin-e and steam from line 26 is contacted with reformer catalyst in reformer 21 likewise operating at about 1500 to 1600 F. and the resulting products, after cooling and separating water, are then passed by compressor 28y to CO2 labsorber 29 into which lean MEA solution is introduced by line 30'. Enriched MEA solution from -the `bottom of tower 29 passes by line 3l and line 20 to CO2 activator tower 2 I- so that the carbon dioxide inthis stream supplements the carbon dioxide from the stream leaving absorber I8 to supply the required amount of CO2 for reformer' 2l'. By the above procedure approximately 1,200,000 cubic feet per day of an approximately 1:1-gas mixture of H2: CO is produced Vand discharged through line 32 while .about 950,000 cubic feet per Aday of a gas consistingl essentially ofuhydrogen but containing.

about .2 tov3=% or more of methane and less than 0.1% of carbon monoxide., is produced and discharged' through line I9.

While. thev invention is applicable to the production of higher alcohols generally, i. e. alcohols containing from about 4 to 16 or more carbon atoms vper molecule, it will be described in Fig. 1 as applied to the production. of nonyl alcohol. In this case a mixture of octenes is obtained by dimerization of oleflns contained in a butanebutylene stream by means of a polymerization catalyst such as phosphoric acid on kieselguhr. The resulting octenes are separated from higher boiling materials by fractional distillation. About 2050 gallons per hour (measured at 60 F.) of the butylene dimer-codimer charge is introduced by line 33, the major part of the charge being pumped to a pressure of 3000 p. s. i. g., preheated at about 330 F. and introduced by line 34 to oxolation `reactor 35 and a minor portion (about gallons per hour) being employed for catalyst recovery as will hereinafter be described.

While a variety of catalysts are known to effect the oxolation reaction, the preferred catalyst in this case is cobalt which is introduced in the form of cobalt naphthenate but which evidently functions as a cobalt carbonyl.y Makeup catalyst may be introduced through line 36 as a 6% cobalt naphthen'ate solution but most of the catalyst is'recovered from oxolation products and recycled as a cobalt naphthenate solution as will be hereinafter described. About gallons per hour of total catalyst solution is introduced into the oxolation reactor which corresponds to about .11 weight per cent cobalt based on olefinl charge. The fresh hydrogen-carbon monoxide gas stream is introduced into the oxolation reactor by compressor 31. oxolation is effected at a pressure of about 3000 p. s. i. g., at a temperature of about 330 F. and a liquid space velocity of about .4 to 1 in unpacked tower reactor 35 which may be about 21/2 to 3 feet in in diameter by about 40 feet in height. Other oxolation temperatures of the order of 250 to 450 F. may be employed with space velocities of the order of about 1. The eiiluent liquid product stream from the oxolation reactor may consist of about 60% unre- `acted oleiins, about 24% of nonyl aldehydes, '7%

nonyl alcohols and 9% higher boiling materials although With proper catalyst and operating conditions the amount of aldehydes may be increased and the amount of high boiling material decreased. The large mass of liquid in the oxolation reactor serves as a temperature stabilizer and minimizes any tendency toward sudden increase in temperature.

The products from the oxolation reactor pass through cooler 38 to separator 39 which may operate at about F. and oxolation pressure and from which the separated gases may be vented through line d0. A part of the separated gases may be recycled through line el by compressor Ha to the base of the oxolation reactor. Sufficient of the cooled liquid products may be recycled by line t2 and pump 42a through spaced points in the reactor in amounts suncient to control the temperature therein. Temperature control is, of course, essential for preventing undesired side reactions and any known means of temperature controlv may be employed.

Liquids from separator 3S are passed through a pressure reducing valve 4S or a throttle system to a low pressure separator lili, which may operate at about li0 p. s. i. g., from which gases are discharged through line 45. Most of such gases may be recycled by line 45- to the inlet of compressor 28 and in admixture with gases from reformer 21 they are scrubbed in absorber 29 to remove CO2. To prevent any build-up of methane, a small amount of gas may be vented continuously or intermittently through line 45a. The recycle of released gas by line 45 slightly increases the effective capacity of the plant.

Liquid from separator 44 is then Washed with about to 10% sulfuric acid which may be introduced at the rate of about 60 to 65 gallons per hour from line 46. The Washing may be effected in one compartment of a horizontal Wash drum 41 provided with a suitable stirrer, the other portion of the Wash drum serving as a separator or settler. Any effective contacting and separating system may be used for this purpose, either in single stage or multiple stage. About 480 gallons per 'hour of the settled `acid may be recycled by pump 48, the remaining approximately 65 gallons per hour of cobalt sulfate in acid solution being introduced by line 49 to cobalt recovery system 50 into which about 75 gallons per hour of a liquid hydrocarbon (olen charging stock) is introduced by line 5|, about 40 gallons per hour of about caustic solution is introduced by line 52 and about 15 to 20 gallons per hour of naphthenic acid is introduced by line 53. The cobalt recovery system may comprise a simple cylindrical vessel provided with a stirrer, the net reaction being a conversion of the cobalt sulfate to cobalt naphthenate, which reaction proceeds almost quantitatively because the precipitated cobalt naphthenate is dissolved in the introduced hydrocarbon. The total mixture from this vessel is introduced by pump 54 into separator 55 from which sodium sulfate solution is withdrawn by line 56. The cobalt naphthenate solution may be Washed with Water in further mixing and separating zones (not shown) to remove all sodium sulfate and any excess caustic. The solution is returned to the oxolation reactor by pump 51 and line 34 together with any make-up cobalt naphthenate that may be required.

The acid-Washed product from tank 41 is introduced into water Wash vessel 58 wherein it is Washed with water introduced through line 59, the Wash Water being withdrawn by line 60. A plurality of washing steps may of course be employed and a small amount of alkali metal hydroxide may be added to the Wash Water to insure acid removal. Gas may be vented through line 58a.

The washed products of oxolation are then introduced by pump 6| through preheaters (not shown) and by pump 62 to first hydrogenation reactor 63 into which hydrogen is introduced from line 64. Hydrogenation may be effected in a single reactor or in a plurality of reactors connected in parallel and it is effected by trickling the liquid over a, bed or beds of hydrogenation catalyst such as copper chromite or cobalt supported on a carrier such as kieselguhr, pumice, alumina, silica gel, alumina gel, porcelain beads, ltros or the like. The hydrogen employed may be that which has previously vbeen utilizedin a subsequent hydrogenation step.

This rst 'hydrogenation step may be effected at about 500 to 3000, e. g. 850 p. s. i. g. and at a temperature of about 350 to 550 F., e. g. about 450 F., under such conditions that most of the aldehydes are converted into alcohols and a minimum amount of the alcohols, preferably less than 10%, are converted into hydrocarbons. The hydrogenation may be so vcontrolled as to avoid saturation of more than about 15% of the olens present but the system is designed to remove the heat of hydrogenation even if all of the olens are saturated. The liquid space velocity Will depend somewhat upon the catalyst employed; with a cobalt-on-pumice catalyst containing about 3% to 10% cobalt space velocities should bein the range of about .1 to .4 based on incoming oxolated liquid or about .2 to 1.2 based on total volumes of liquid charged per hour per volume of catalyst space. More than of the aldehyde may be converted to alcohol in this step.

The liquid leaving the base of the hydrogenation reactor or reactors may be at a temperature of about 480 to 500 F. or more due to the exothermic nature of the hydrogenation. The hot liquid passes through cooler 65a and pressure release valve 66 to recycle separator 61 which may be operated at about 335 p. s. i. g. and approximately 450 F. About 5000 gallons per hour of liquid from the base of the separator is recycled by line 68 for admixture with the approximately 2200 gallons per hour of washed charged from pump 6l and the preheaters. By thus using a recycle ratio of about 2:1 to 3:1 the temperature rise in reactor 63 may be minimized. The cooling and separation at lower pressure is essential for proper operation of pump 62. Liquids and gases from the upper part of separator 61 are Withdrawn through cooler 69 to separator 10 which may operate at about 335 p. s. i. g. and about F. and from which additional gas is withdrawn through line 1l. The 600,000 cubic feet or more per day of hydrogen withdrawn at this point may be employed in other re-nery units such as hydrogenation, hydroforming of coke still naphtha, desulfurization over cobalt molybdate catalyst, etc. Where a source of carbon monoxide is available this gas may be admixed with carbon monoxide and employed in the oxolation step.

From separator 10 the product stream passes through pressure reducing valve 12 to low pressure separator 13 which operates at about 30 to 40 p. s. i. g. and from which additional hydrogen is rvented through line 14. The low temperature and low pressure gas separation is foreffecting substantially complete degasication and for avoiding loss of products with vented gas. Gases may be vented at higher temperature with a condenser for preventing loss of product, but the above system provides a more complete removal of hydrogen.

The separated product, approximately 2400 to 3000 gallons per hour, then passes through heater 15 and line 16 to fractionator 11 which is provided with reboiler 18 at its base. The fractionator is preferably operated at subatmospheric pressure, i. e. about 10 p. s. i. a. with a bottom temperature of about 375 F. and a top temperature of about 220 F'. The overhead is cooled in condenser 19 and introduced into receiver 8i) from which gases and vapors are discharged through line 8l by means of an ejector for maintaining the partial vacuum. Part of the liquid from receiver 80 is recycled as reflux through line 82 and the remainder withdrawn through line 83. About 1200 gallons per hour of codimer gasoline is thus withdrawn from the system.

It has been found that a side stream withdrawn from fractionator 11 between the said inlet and the top of the tower is richer in aldehydes than a stream which is being introduced into the fractionator. About 10 to 35% of the oxolation product or, in other Words, about 200 to 700 gallons per hour of such side stream .is vtherefore Withdrawn through one of the branched trap-'out lines 8d and returned by pump 85 .and line 85a to hydrogenation reactor 63. The recycle of this particular materialto the hydrogenation reactor oiiers many advantages. It decreases the amount of aldehydes removed with .codimer from the top of tower 'il and the amount of aldehydes which must be hydrogenated :in the second hydrogenation step. eiectiveness of the first hydrogenation reactor by decreasing the relative amount of high boiling material which would have to be returned thereto vif the total recycle was via line 08. It greatly yreduces .the amount of hydrogenation that must be effected in the final step and in some cases may even avoid the necessity of vsuch final hydrogenation step.

vThe alcohol and high boiling .components are removed from the bottom of fractionator 1.1 and introduced to rerun tower 80 which is provided with rcboiler Si and stripping gas inlet 8.8 .and 'from which high boiling components are Withdrawn through line 89. Stripping shouldy .be e'iected with an inert gas 'such' for example as hydrogen. The nonyl alcohol `fraction is condensed in cooler and introduced into rreceiver Sil from which a part is returned as reiiux through line 02 and-another part is pumped to a pressure of about 900 p. s. i. g. by Vpump 03 and introduced by line 0 to second hydrogenation reactor 05. The catalyst in this second reactor may be the sanie as in reactor t3 and this second hydrogenation is preferably effected at a pressure of about 500 to 3000, e. g. 900 p. s. i. g., with an linlet temperature of about 400 F. and .an outlet temperature of about `VLi F., .the liquid rspace velocity being about the same yas that employed in the iirst hydrogenation, namely, .about .2 to v1.0-volume of liquid perhour per volume of catalystspace. The `hydrogen .is introduced into 'reactor .Baby line I9 and compressor 00, the vunused .and liuidissolved hydrogen being vented lfrom the base of the reactor through line at labout 410 for use in the rst hydrogenation reactor 53. Approximately '7100 .gallons perhour of the impure nonyl alcohol is thus introduced linto reactor 95 and approximately 50 to .60 cubic feet of hydrogen is introduced thereto per gallon of impure alcohol to be treated.

The hydrogenated product is cooledin cooler al' and introduced through'pressure 'reducing valve "Sli to separator 90 from which hydrogen yis vented through line |100 and the final -nonyl alcohol is withdrawn through line lill. IIlhe nonyl alcohol thus produced is usually of market- "able grade withoutany further treatment but the hydrogenation in reactor 95 is of such kseverity as to reduce lany alcohols to lhydrocartbons', a Vsubsequent! distillation step maybe 're- D"quired, -The non-yl alcohol produced as 'herein-- above described may have a refractive index (comici 1.433, a specic gravity (15.'56/15.561C`.) -oi about :843, an aldehyde content less than .5%,

ra neutral-ization number of about 02 milligram ou.

oi KOH per gram, a color yof about Saybolt and a-fdistil'lation range of about 193 to 196 C. and an ASTM distillation range (10% to 90% points) uconsists essentially .of C5 to C12 olens'and 'boils -in thema-nge of about `100 to 'about 400 F. sSuch feed may vberirst reedfromCa 'and heavierolens It increases the 10 in :a rst'tower and then freed fromCe `and lighter olens in a second tower. In this example such feed is .introduced by ,line 3 into fractionator .4 which is provided with conventional reboiler and reflux means and which .is operated to remove the C5 and Ct oleiins as an overhead vapor stream to gasoline line 6. Vlhe bottoms from fractionator tl are introduced by line l into fractionator e rroinwhich Cs and higher oleiins are withdrawn bottoms through line 9 and added to gasoline line v6. The Cn olefins Withdrawn from rthe top of iractionator 8 preferably contain not more than 1% of C6 voleins and not more than 2% .of Ccolerins. .Aboute barrels per day of such C7 Yolefin stream is introduced by line 33 to oxola.-

tionreactorii together withabout 430,000 cubic feet per day of hydrogen and carbon monoxide (in a mol ratio of 12.1 to 1.3:1) throughline .32 and about 31/2 barrels per day :of cobalt catalyst solution through. line 3ft. The catalyst in this case is a .6% cobalt .naphthenate solution in C1 oleiin .employed in such amounts as to provide about .1 weight percent of .cobalt basedxon the total C1 voleiins charged, but other oilsoluble cobalt `saltsmay be employed.

The -oxolation .reactor is operated :at 30.00 p. s. `i. g., .about 325 F'. anduabout 5 space velocity `to obtain Va %k conversion of the olen feed, '8% to 10% `of said feed being reduced to Cvparaflin, about'30% being converted to Cs aldehyde, about 8 to 10% to C8 .alcohol and about 12% to polymer. The temperature .control is effected as described in connection with Fig. 1, about 3 .or d volumes oi' product liquid being recycled for each volume of olein feed introduced and gas vbei-ng recycled to the bottom of .the reactor in amounts necessary for .temperature control at that point. More specifically, the total .eluent stream which leaves the top of oxolation reactor 35 is. .cooled at reaction pressure in cooler 38 and the .cooled products are introduced into receiverv 3S which .is'hcld at. about 100 F. Uncondensed. gases leave the top .of receiver 3:9', a part thereof being vented through line .40 and the remainder .of

these .cooled gases being recycled .by compressor Ma to 'the inlet end of the voxolation reactor below the level at Which portions of condensed cool lliquid are returned through line d2 .by pump Valla. Thus the cooled gas recycled through line AI traverses the entire reactor While the cooled liquid from line 132 is introduced r`only at upper .levels in the reactor.

The net prod-uct liquid from highp'ressure sep- .aratortd passes through release valvedt .to 210W pressure separator fil which .operates at yabout 30 p. s. i. g. and about 100 F., .liberated gas being vented .to a fuel line through line '.45'.

The product liquid, about 300 :barrels -lOer .day

Acontaining approximately l111.4 mols per hour vof valdeyhyde and alcohol .are 'adinixed with about l10% sulfuric acid from .line et" `and .intimately mixed therewith with approximately one hour holding 'time in *mixer LM. Ii desired, .separator dit", may-.be omitted, lthe .stream passing through Iline tia .to di and gas 'being vented through line'i'lfiE-a. .Aipartiof .the separatedacid maybe continuously recycled by pumpall while another part is .continuously lwithdrawn through -line lfil 'for I.recovery of thecataly-.st component. Due to the reaction of :about half or ,the acid .to form CoSOi, theactual acid strength in the :acid washing and .separation stage may be only about 5%.`

As :in the previous example, theextract'with drawn through line de' is `introduced .to vcobalt :through line 56' ..cobalt salt is returned by pump 5l to line 3d' .tion is about 3 feet 6 inches by 11 feet.

vrecovery system 50', the cobalt being present in the extract as an aqueous solution of cobalt sulfate regardless of the particular oil-soluble colbalt salt which is employed in the catalyst solution. While cobalt naphthenate is the preferred cobalt soluble salt, we may employ cobalt tallate (the cobalt salt of acids contained in the tall oil by-product obtained in paper manufacture), cobalt stearate, cobalt oleate or any other cobalt salt which is soluble inthe hydrocarbon charged to the oxolation reactor. By contacting the extract from line 49 with an alkali metal hydroxide solution, such as sodium hydroxide, potassium hydroxide, lithium hydroxide, or any equivalent thereof, in the presence of an oil-soluble acid, such as naphthenic acid, tallic acids, stearic acid, or a preferentially oil-soluble carboxylic acid, and also in ther presence of a hydrocarbon diluent, such as a part of the olefin feed to the oxolation reactor, Water-soluble alkali metal sulfates are obtained and the cobalt is recombined with the Voil-soluble acid which in turn is dissolved in the hydrocarbon diluent. The aqueous alkali metal sulfate can then be separated and withdrawn while the diluted oil-soluble for further use in the process, any necessary make-up catalyst being introduced by line 36. After catalyst removal from the oxolation ,product stream by acid wash, said stream is washed in mixer 58 with water introduced by line 59', the waste water being discarded through line 60'. It is important in this case not only that catalyst be eliminated from the loxolation product stream but also that the product stream then be freed from acid before it is subjected to further treatment. A small amount of alkali metal hydroxide may be employed in the wash water to insure a pH of `about 7 in the washed stream.

After acid washing to remove cobalt catalyst .and water Washing to remove acid, the oxolation effluent stream is passed by line |02 with steam from line |03 and/or with recycled oleiins through exchanger |04 where it is heated to about 200 F., and thence to an intermediate level in fractionator |05 which in this case is operated at a pressure of about 180 to 220 mm. of mercury. That portion of fractionator |05 vwhich is above the feed inlet is preferably of larger diameter than the portion below the feed inlet since it is desirable to effect as much flash distillation as possible. Thus the fractionator portion of the tower may be about 6 feet in diameter by 31 feet tall, while the stripping sec- Stripping steam may be introduced at the base of the narrowed section of tower |05 through line |06. The tower should be operated with as low a temperature and pressure and with as short a contact time as is economically feasible and the tower bottom temperature in this case does not exceed about 250 F. Somewhat higher preheat and bottom temperature may be required when making higher boiling aldehydes and alcohols, but it is preferred to keep. below 350 F.

Components lower boiling than Cs aldehydes (chiefly unconverted C7 olens and C7 olens which have been saturated) are withdrawn from -the top of the tower through cooler |01 to receiver |0711 from which about 120 barrels per day is passed by line |011) to gasoline line 6,

about 510 barrels per day is returned to the -top of tower |05 to serve as reflux and about Aif) 1340 to1350 barrels per day is recycled by line |07c for admixture with steams being introduced by lines |02 and |03 to preheater |04. The recycle of about 4 or 5 volumes of C7 hydrocarbons per volume of fresh feed to the fractionator-preheater makes it possible to limit the tower inlet temperateure to about 200 F. and the tower bottom temperature to 250 F. and to obtain effective flash vaporization of the Ca aldehyde-alcohol components while minimizing the loss of aldehyde by aldol condensation and avoiding the mechanical difculties in the tower which could result from the condensation of water on the trays which are employed therein if steam alone were used to meet these temperature limitations. While the recycle of C1 hydrocarbons is particularly advantageous for providing a combined heat carrier and stripping medium, other known methods may be employed for effecting flash distillation-stripping operation provided that the temperature requirements are met so that substantially all materials higher boiling than C'a alcohol can be withdrawn from the bottom of tower |05 through line |00 and so that the stream which is Withdrawn from a trap-out plate above feed inlet through line |00 will consist essentially of Ca aldehyde and alcohol which is substantially free from higher boiling materials. From the standpoint of product quality, it is not essential to remove all components lower boiling than Ca aldehyde. If all of the lower boiling components are retained with the Ca aldehyde-alcohol fraction, the load on the subsequent hydrogenation tower may be increased to an undesirable extent and in this example, the heart cut withdrawn through line |09 may contain as much as 10 or even 20 mol per cent of C7 hydrocarbons (chiefly olefins).

By the intermediate fractionation step about barrels per day of Cs aldehyde-alcohol (containing about 10 to 20% C7 hydrocarbon) is withdrawn for hydrogenation. This fraction is heated in reactor furnace I0 together with make-up and recycled hydrogen from line |09a and introduced into reactor which in this case is operated at a pressure of about 3000 p. s. i. g. and at a temperature in the range of 350 to 600 F. The catalyst is preferably copper chromite or about 12% cobalt on pumice. With a fresh feed liquid space velocity of about .25 the oleflns are completely saturated and the hydrogenation of the -pressed by compressor ||1 and returned by line ||8, manifold ||9 and spaced inlets |20 to prevent a temperature rise of more than 25 F. in

any part of the reactor. About one volume of -the cooled liquid product withdrawn through line |2| is recycled by pump |22 and line |23 for each volume of heart cut fraction introduced by line |09. In this case about 16 mols per hour of fresh hydrogen is introduced by line |09a.

The net hydrogenation product passes through pressure reducing valve |24 and is introduced into low pressure separator |25 which is maintained at about 30 p. s. i. g. and at about 125 F. Liquid from this separator is introduced through line |21 to fractionator |28 which is provided with usual reboiler and reflux means and materials boiling lower than oxo alcohol (chiey hydrogenated olens) are .taken overhead through line v|29 to gasoline line 6. Bottoms from tower |28 are introduced Yby line |30 into final fractionator I3| for removing oxo bottoms through line |32 from the final oxo alcohol (octyl alcohol in this case) which is taken` overhead through line |33. The iso-octyl alcohol thus produced boils 'from about 360 to 370 F., has a flash point of about 170 to 180 F., is Water White, has an aldehyde content which is less than about 2%., may contain a slight amount of iso-heptyl alcohol and iso-nonyl alcohol, depending upon the `efficiency of the fractionation of the original ole- -fin feed, but the iso-heptyl content usually does not exceed .2% and the iso-nonyl alcohol usually does not exceed 2.0%.

While we have described in detail specific examples of our invention, it should be understood that various alternative operating procedures and operating conditions will be apparent from the above description to those skilled in the art.

With regard to the hydrogenation catalyst it should' be understood that the invention is not limited to cobalt-on-pumice nor even to supported cobalt since nickel, copper chromite, cobalt molybdate and other known hydrogenation catalysts may be employed. Preferably the unreduced hydrogenation catalyst is charged to the reactor for activation by reduction in situ. The catalyst is heated to a vtemperature within the range of Aabout 525 F. to about 650 F. by means of superheated steam. During this preheat step the pressure in the reactor is raised to a range of about 300 to 450 pounds per square inch to obtainv maximum rate of heat input. After the catalyst has been heated to the `desired temperature the reactor pressure is reduced to atmospheric and the activationstep is carried out with hydrogen. Alternatively the heating of the catalyst to reduction temperature may be effected by simply circulating hydrogen, diluted with an inert gas if desired, through reactor heater H0, thence through the unreduced .catalyst in reactor IH and thence to a condenser-separator (not shown) for removing the Water formed in the reducing step.V Completion of activation is marked by the absence of water vapor in the efiiuent gas from the activation step.

With regard to the olefin charged, we may employ Cs clef-ins obtained from catalytically cracked naphtha, from hydrocarbonsynthesis r from any other hydrocarbon; in this case the hydrocarbon fraction containing the C8 olefins may be charged to the oxolation reactor and the space velocity modified on account of the large amount of. inert paraflins which would be present. Similarly, anarrow cutCs, Cs, Cfz, C12 or in fact any particular narrow cut olefin fraction may be employed, the yield of alcohol in each case depen-ding upon the amount and nature of the olefin present and the operating conditions employed. For oxolation the olen must have at least 1 carbon atom attached both to a double bond and to at least 1 hydrogen atom; however, since olens which are not oxolatable may be isomerized under oxolation conditions into oxolatable form, the invention is applicable tor any aliphatic yolefin having at least three and preferably about four to fifteen carbon atoms per molecule.

An. important feature of the invention is the production of alcohols which are substantially free from color bodies and color forming materials. An important use ofthe alcohols is the preparation therefrom of esters such for example as dioctyl (or dinonyl phthalatewh'ich is a "14 valuable plastici'zer; Not only must the alcohol be free from color bodies but it must also be free from materials which cause color formation or ph'thalation. The APHA color of the ester should not exceed 200 and it should preferably be less than in the'case of octyl alcohol.

All volumes referred to herein are those measured at 60 F. and atmospheric pressure.

We claim:

1. The method of making alcohols which `coinprises reacting an olefin having in the range of 3 to 15 carbon atoms per molecule with a` carbon monoxide-hydrogen gas having a m01 ratio of about 1:1 in the -presence of an oil-soluble cobalt salt dissolved in a hydrocarbon solvent in a reaction zone under pressure of about 2500 to 4000 pounds per square inch at a temperature of about 250 to 400 F., cooling the elliuent stream from said reaction zone and separating gas from the cooled stream, recycling yat least a portion of said cooled stream at spaced points above the bottom of the reaction zone and recycling at least a portion of said separated gas to the bottom of said reaction zone for removing exothermic heat of reaction therein, washing cobalt from thev remaining liquid stream with sulfuric acid to form a cobalt sulfate extract, separating said 'extract from said product stream, reacting said extract with an alkali metal hydroxide solution and an oil-soluble acid in the presence of a liquid hydrocarbon whereby an oil-soluble cobalt salt is reformed and dissolved in said hydrocarbon as a catalyst solution and alkali metal sulfate is formed, separating said alkali metal sulfate and any unreacted alkali metal hydroxide from the catalyst solution, returning said catalyst solution to said reaction zone, and hydrogenatin'g at least a portion of the acid-Washed stream to convert aldehydes contained therein to alcohols.

2. The method of claim l which includes the further step of 'Washing the product stream with water 'after it is Washed With sulfuric acid.

3. The method of claim 2 which includes the steps of fractionatin'g' at subatinospheric pressure the acid-washed and water-washed oxolation products which have been partially hydrogenated and subsequently hydrogenating a. 'fraction obtained in the fractionating step which fraction is substantially free from compounds having a higher number of carbon atoms than theA desired alcohol.

4. The method of claim 1 which includes the further step of was-hing sodium sulfate and any unreacted sodium hydroxide from the catalyst solution with water before said solution is returned to the reaction zone.

5. The method of claim l wherein the oilsoluble cobalt salt is cobalt naphthenate and the oil-soluble acid is naphthenic acid.

6. In the process of producing alcohols by reacting aliphatic olefins having at least three carbon atoms per molecule with a carbonxmonoxidehydrogen gas having a mol ratio of about 1:1 in the presence of a cobalt catalyst in ay reaction zone under conditions for converting a substantial portion but not fall' ofthe olens into aldehydes having one more carbon atom: per molecule than the olens and subsequently converting most of said aldehyd'es 'to alcohols in the presence of unreacted iolens in a liquid stream, the infrproved method of operation which comprises effecting reaction betweenv said olefins and carbon monoxide-hydrogen gas in the presenceof.v a cobalt catalyst supplied by adding anoi'l-soluble cobalt salt tosaid reaction: zone, vseparating gases `about 250 to 400 F., whereby at least peri 15 fromreactio-n zone eiuent liquid, Washing cobalt from theremaining eiliuent liquid with sulfuric 'acid to form a cobalt sulfate extract, separating said extract from said eiiluent liquid, reacting said extract with a sodium hydroxide solution and an oil-soluble acid in the presence of a liquid hydrocarbon whereby an oil-soluble cobalt salt is reformed and dissolved in said liquid hydrocarbon, returning said hydrocarbon and dissolved oil-soluble cobalt salt to the reaction zone, effecting conversion of aldehydes in the effluent liquid, which has been freed from catalyst by sulfuric acid washing, to alcohols by hydrogenation in the presence of a supported cobalt hydrogenating catalyst at a pressure in the range of 500 to 3000 pounds per square inch at a temperature in the range'of about 400 to 550 F., cooling the liquid leaving the hydrogenation step and reducing the :pressure on said liquid for effecting separation of hydrogen from said liquids, reheating the liquid after hydrogen removal, fractionating the reheated liquid to. separate a low boiling olefin fraction, a higher boiling aldehyde fraction, a still higher boiling alcohol fraction, and a highest boiling fraction and recycling said higher boiling aldehyde fraction to said hydrogenation step.

7. In the process of producing alcohols by reacting aliphatic olens having at least three carbon atoms per molecule with a carbon monoxidehydrogen gas having a mol ratio of about 1:1 in the presence of a cobalt catalyst under conditions for converting a substantial portion but not all of the oleflns into aldehydes having one more carbonatom per molecule than the olens and subsequently converting most of said aldehydes to alcohols in the presence of unreacted olefins in a liquid stream, the improved method of operation which comprises washing said liquid stream with dilute sulfuric acid to remove metallic cobalt and cobalt compounds therefrom, contacting said acid washed stream with a supported hydrogenation catalyst under a pressure in the range of 500 to 3000 p. s. i. and yat a temperature in the range of about 400 to 550 F. fo-r converting most of the aldehydes to alcohols without saturating'most of the olens nor converting as much as of the vproduct alcohols into hydrocarbons, depressuring the hydrogenated stream and removing gas therefrom, then fractionating said stream to obtain an olefin fraction, an alcohol fraction containing some aldehyde, an aldehyde fraction higher boiling than the olefin fraction and lower boiling than the alcohol fraction, and a high boiling fraction, removing the aldehyde content of said alcohol fraction, and recycling said aldehyde fraction to said hydrogenation step.

8. The method of producing an alcohol which Ycomprises continuously introducing at the base of a vertical reaction zone a fresh feed liquid -consisting essentially of an olefin having at least three carbon atoms per molecule and for each gallon of fresh feed about 20 to 50 cubic feet of fresh hydrogen-carbon monoxide gas mixture having a mol ratio of about 1:1 together with a sufficient amount of oil-soluble cobalt salt dissolved in a portion of the fresh feed to. provide an amount of cobalt in the range of .0l to .2 per cent based on fresh feed, passing said mixture upwardly through said reaction Zone at a space velocity in the range of .15 to 1.5 Vo-lumes of fresh liquid feed per hour per volume of reactor space under a pressure in the range of about 2500 tr;

at a temperature in the range o 4000 p. s l g about 30 cent of said olefin is converted into oxygen- 'ated compounds containing one more carbon atom than said olefin and about one-fifth to onehalf of said compounds are alcohols formed by initial hydrogenation and the rest `are aldehydes, higher boiling materials being simultaneously produced, cooling the effluent stream from the reaction zone and separating the stream into a cooled gas stream and a cooled liquid product stream, recycling a part of said cooled gas stream at a point near the base of said reactionzone and recycling cooled liquid product at spaced higher points in the reaction zone in amounts sufficient to contro-l temperatures therein, reducing thepressure on the unrecycled liquid product stream, removing gases liberated on reduction of pressure, washing said stream with dilute sulfuric acid to remove cobalt therefrom as cobalt sulfate, freeing said stream from acid, subsequently fractionating said stream at subatmospheric pressure toremove therefrom substantially all materials higher boiling than the alcohol component of said compounds, hydrogenating at least' the aldehyde component o-f the remaining stream by final hydrogenation in the absence of said higher boiling materials by contacting said aldehyde component with a hydrogen-ating catalyst at a pressure in the range o-f about 500 to about 3000 p. s. i. g. with a space velocity in the range of about .15 to 1.5 volumes of fresh liquid charged to the hydrogenation step per hour per volume of catalyst space, at a temperature in the range o-f 350 to 550 F. with an excess of hydrogen substantially free from carbon monoxide whereby substantially all of the aldehyde compounds are converted to alcohols and only a small amount of saturated hydrocarbons are formed and fractionating the products of the final hydrogenation to recover therefrom an alcohol having one more carbon atom per mo-lecule than the original olefin charged, which alcohol is substantially free from hydrocarbons, aldehydes and color-forming materials.

9. The method of claim 8 which includes the step of removing from the stream in the subatmospheric fractionating step most components Alower boiling than the aldehyde component of said compounds as well as substantially all components higher boiling'than the alcohol component of said compounds.

LEONARD W. RUSSUM. ROBERT' J. HENGSTECBECK.

References cited in the sie of this patent UNITED STATES PATENTS Number Name Date 2,437,600 Gresham et al. Mar. 9, 1948 2,464,916 Adams etal Mar. 22, 1949 2,504,682 Harlan Apr. 18, 1950 2,530,989 Parker Nov. 21, 1950 l 2,557,701 Smith et al June 19, 1951 OTHER REFERENCES IGF Patent Applications T. O'. M. Reel 36, Item 21 and part of Item 36, Application 172,948 IVd/ 120, O. Z. 13,599, August 10, 1942. Deposited (Also available in Meyer translation pp. 47-49.) 

1. THE METHOD OF MAKING ALCOHOLS WHICH COMPRISES REACTING AN OLEFIN HAVING IN THE RANGE OF 3 TO 15 CARBON ATOMS PER MOLECULE WITH A CARBON MONOXIDE-HYDROGEN GAS HAVING A MOL RATIO OF ABOUT 1:1 IN THE PRESENCE OF AN OIL-SOLUBLE COBALT SALT DISSOLVED IN A HYDROCARBON SOLVENT IN A REACTION ZONE UNDER PRESSURE OF ABOUT 2500 TO 4000 POUNDS PER SQUARE INCH AT A TEMPERATURE OF ABOUT 250* TO 400* F., COOLING THE EFFLUENT STREAM FROM SAID REACTION ZONE AND SEPARATING GAS FROM THE COOLED STREAM, RECYCLING AT LEAST A PORTION OF SAID COOLED STREAM AT SPACED POINTS ABOVE THE BOTTOM OF THE REACTION ZONE AND RECYCLING AT LEAST A PORTION OF SAID SEPARATED GAS TO THE BOTTOM OF SAID REACTION ZONE FOR REMOVING EXOTHERMIC HEAT OF REACTION THEREIN, WASHING COLBALT FORM THE REMAINING LIQUID STREAM WITH SULFURIC ACID TO FORM A COLBALT SULFATE EXTRACT, SEPARATING SAID EXTRACT FROM SAID PRODUCT STREAM, REACTING SAID EXTRACT WITH AN ALKALI METAL HYDROXIDE SOLUTION AND AN OIL-SOLUBLE ACID IN THE PRESENCE OF A LIQUID HYDROCARBON WHEREBY AN OIL-SOLUBLE COBALT SALT IS REFORMED AND DISSOLVED IN SAID HYDROCARBON AS A CATALYST SOLUTION AND ALKALI METAL SULFATE IS FORMED, SEPARATING SAID ALKALI METAL SULFATE AND ANY UNREACTED ALKALI METAL HYDROXIDE FROM THE CATALYST SOLUTION, RETURNING SAID CATALYST SOLUTION TO SAID REACTION ZONE, AND HYDROGENATING AT LEAST A PORTION OF THE ACID-WASHED STREAM TO CONVERT ALDEHYDES CONTAINED THEREIN TO ALCOHOLS. 